Production of titania

ABSTRACT

A sulfate process for producing titania from a titaniferous material is disclosed. The process includes leaching the titaniferous material and producing a leach liquor, separating titanyl sulfate from leach liquor, hydrolysis of the extracted titanyl sulfate, and thereafter calcining the solid phase produced in the hydrolysis step. The process is characterised by multiple stage leaching of the titaniferous material.

The present invention relates to a process for producing titania from atitaniferous material.

The term “titaniferous” material is understood herein to mean anytitanium-containing material, including by way of example ores, oreconcentrates, and titaniferous slags.

The present invention relates particularly to the sulfate process forproducing titania from titaniferous material.

The sulfate process was the first commercial process for the manufactureof titania from titaniferous ores, such as ilmenite.

A significant issue with the sulfate process is that it produces largequantities of waste iron sulfate and consumes large quantities ofsulfuric acid.

The chloride process generally avoids the iron sulfate waste problem ofthe sulfate process and, at larger scales, is less expensive to operatethan the sulfate process.

Hence, the chloride process is the currently preferred process forproducing titania, particularly titania for the pigment industry.

An object of the present invention is to provide an improved sulfateprocess.

In general terms, the present invention provides a sulfate process forproducing titania from a titaniferous material (such as ilmenite) whichincludes the steps of:

(a) leaching the solid titaniferous material with a leach solutioncontaining sulfuric acid and forming a leach liquor that includes anacidic solution of titanyl sulfate (TiOSO₄) and iron sulfate (FeSO₄);

(b) separating the leach liquor and a residual solid phase from theleach step (a);

(c) separating titanyl sulfate from the leach liquor from step (b);

(d) hydrolysing the separated titanyl sulfate and forming a solid phasecontaining hydrated titanium oxides;

(e) separating the solid phase containing hydrated titanium oxides and aliquid phase that are produced in the hydrolysis step (d); and

(f) calcining the solid phase from step (e) and forming titania;

and is characterised by the steps of:

(i) a further leach step of leaching the residual solid phase from step(b) with a leach solution containing sulfuric acid and forming a leachliquor that includes an acidic solution of titanyl sulfate and ironsulfate and a residual solid phase;

(ii) separating the leach liquor and the residual solid phase from step(i); and

(iii) supplying the separated leach liquor to the leach step (a) and/ormixing the separated leach liquor with the leach liquor from step (b).

The term “hydrated titanium oxides” is understood herein to include, byway of example, compounds that have the formula TiO₂.2H₂O and TiO₂.H₂O.

In addition, the term “hydrated titanium oxides” is understood herein toinclude compounds that are described in technical literature as titaniumhydroxide (Ti(OH)₄).

The leach step (a) and the further leach step (i) may be carried out inthe same vessel.

In that event, the further leach step (i) includes returning theresidual solid phase from step (b) to the vessel, wherein the residualsolid phase forms part of the titaniferous material subjected toleaching in step (a).

Alternatively, the leach step (a) and the further leach step (i) may becarried out in separate, with the residual solid phase from step (b)being supplied to the vessel or vessels for the further leach step (i).

In that event, preferably the process includes separating the leachliquor and a further residual solid phase formed in the further leachstep (i).

The separated leach liquor may be supplied to the leach step (a).

Alternatively, the separated leach liquor may be mixed with the leachliquor from step (b) and thereafter the mixed leach liquor may beprocessed in the subsequent steps of the process.

The leach step (a) and/or the further leach step may be carried out on acontinuous basis or a batch basis.

The applicant has found in experimental work that it is important tocarry out the leach step (a) and/or the further leach step (i) underleach conditions, described herein, that avoid an undesirable amount ofpremature hydrolysis of hydrated titanium oxides.

In addition, the applicant has found in experimental work that it isimportant to carry out the leach step (a) and/or the further leach step(i) under leach conditions that avoid an undesirable amount of prematureprecipitation of titanyl sulfate.

Preferably the leach step (a) and/or the further leach step (i) includeselecting and/or controlling the leach conditions in the leach step (a)and/or the further leach step (i) to avoid undesirable amounts ofpremature hydrolysis of hydrated titanium oxides and undesirable amountsof premature precipitation of titanyl sulfate.

The relevant leach conditions include any one or more than one of acidconcentration, leach temperature and leach time.

Typically, the acid concentration in the leach step (a) and/or thefurther leach step (i) should be at least 350 g/l sulfuric acidthroughout the leach step (a) and/or the further leach step (i) whenoperating at a leach temperature in the range of 95° C. to the boilingpoint in order to avoid premature hydrolysis.

Typically, the acid concentration at the end of the leach step (a)and/or the further leach step (i) should be less than 450 g/l whenoperating at a leach temperature in the range of 95° C. to the boilingpoint in order to avoid an undesirable amount of premature precipitationof titanyl sulfate.

It is noted that the acid concentration at the start of the leach stepcould be higher, typically as high as 700 g/l.

Typically, the leach conditions should be selected and/or controlled sothat the titanium ion concentration in the leach liquor is less than 50g/l in the leach liquor at the end of the leach step (a) and/or thefurther leach step (i).

Preferably the titanium ion concentration in the leach liquor is 40-50g/l.

Preferably the process includes carrying out the leach step (a) in thepresence of an additive that accelerates the rate of leaching thetitaniferous material.

Preferably the process includes carrying out the further leach step (i)in the presence of an additive that accelerates the rate of leaching thetitaniferous material.

The use of the leaching accelerant makes it possible to use lessconcentrated sulfuric acid than is required for the conventional sulfateprocess.

Preferably the leaching accelerant is selected from a group thatincludes iron, a titanium (III) salt, a thiosulfate salt, sulfur dioxideor any other reduced sulfur containing species.

Preferably the process includes carrying out the leach step (a) in thepresence of a reductant that reduces ferric ions to ferrous ions in theacidic solution of titanyl sulfate and iron sulfate produced in theleach step (a).

Preferably the process includes carrying out the further leach step (i)in the presence of a reductant that reduces ferric ions to ferrous ionsin the acidic solution of titanyl sulfate and iron sulfate produced inthe leach step (a).

The reductant may be any suitable reductant.

Preferably the reductant is selected from a group that includes iron, atitanium (III) salt, a thiosulfate salt, sulfur dioxide or any otherreduced sulfur containing species.

As is indicated above, the purpose of the reductant is to minimise theamount of iron in the trivalent ferric form and to maximize the amountof iron in the divalent ferrous form in the leach liquor produced in theleach step (a) and/or the further leach step (i). Maximising the amountof iron in the divalent ferrous form minimises the equilibriumconcentrations of iron in the circuit, by promoting the precipitation offerrous sulfate, for example FeSO₄.7H₂O.

Preferably the process includes the steps of precipitating iron sulfatefrom the leach liquor from step (b) and separating precipitated ironsulfate from the leach liquor prior to the titanyl sulfate separationstep (c).

Preferably the process includes using at least part of the leach liquorremaining after separation of titanyl sulfate in step (c) as at leastpart of the leach solution in the leach step (a) and/or in the furtherleach step (i).

The use of the depleted leach liquor from the titanyl sulfate separationstep (c) as the leach solution for leach step (a) and/or the furtherleach step (i) is an advantage of the process because it maximises theeffective use of acid in the process.

Furthermore, the use of the depleted leach liquor allows a reduction orcomplete elimination of the production of waste acidic effluents and/ortheir neutralisation products, such as “brown gypsum”.

Furthermore, the use of the depleted leach liquor allows recuperation ofheat and also eliminates energy intensive acid recovery and evaporativeconcentration steps.

Preferably the titanyl sulfate separation step (c) includes a solventextraction step of extracting titanyl sulfate from the leach liquor fromstep (b) into a solvent and thereafter stripping titanyl sulfate fromthe solvent and forming a solution that contains titanyl sulfate andthereafter hydrolysing the titanyl sulfate-containing solution in thehydrolysis step (d).

In a situation in which the titanyl sulfate separation step (c) is asolvent extraction step, preferably the process includes using at leastpart of a raffinate from the solvent extraction step as at least part ofthe leach solution in leach step (a) and/or in the further leach step(i).

Preferably the leach solution in the leach step (a) and the furtherleach step (i) includes the raffinate and make-up fresh sulfuric acid.

Preferably the raffinate from the solvent extraction step has an acidconcentration of at least 250 g/l sulfuric acid.

Preferably the raffinate from the solvent extraction step has an acidconcentration of at least 350 g/l sulfuric acid.

Preferably the solvent extraction step includes contacting the leachliquor from step (b) with the solvent and a modifier.

The term “solvent” is understood herein to mean a reagent and a diluentin combination.

The term “modifier” is understood herein to mean a chemical whichchanges the solubilising properties of the solvent such that thetitanium containing species are soluble in the solvent at higherconcentrations than might otherwise be possible.

Preferably the process includes controlling the hydrolysis step (d) toproduce a selected particle size distribution of the hydrated titaniumoxides product.

The controlled growth of coarse particles of hydrated titanium oxides inthe hydrolysis step (d) is a significant departure from the conventionalsulfate process in which there is a strong preference for producing fineparticles in order to produce fine titania that meets the needs of thepigment industry, the major user of titania.

There are some applications, such as electrochemical reduction oftitania, in which it is preferable to have a coarse feed of hydratedtitanium oxides or a coarse feed of titania.

For these applications, preferably the process includes controlling thehydrolysis step (d) to produce coarse hydrated titanium oxides, ieoxides having a particle size of at least 0.005-0.01 mm (ie 5-10micron).

Equally, there are other applications, such as production of pigments,in which it is preferable to have a fine feed of hydrated titaniumoxides or a fine feed of titania.

For these applications, preferably the process includes controlling thehydrolysis step (d) to produce fine hydrated titanium oxides, ie oxideshaving a particle size of less than 0.0003 mm (ie 0.3 micron).

Preferably the process includes using the liquid phase produced inhydrolysis step (d) as a source of acid or water in other steps of theprocess. Typically, the liquid phase includes 100-500 g/l sulfuric acid.By way of example, the liquid phase may be used as a source of acid (andtitanium values) by direct addition to leach liquor, depleted leachliquor or any one of steps (a) and (b) and the further leach step (i).By way of further example, the liquid phase may be used as a source ofwater for washing solid products from any one of steps (b) and (e).

Alternatively, the process may include treating the liquid phaseproduced in hydrolysis step (d) by neutralising the acid in the liquidphase with lime (CaO) and/or limestone (CaCO₃) and producing cleangypsum (CaSO₄.2H₂O).

It is known to produce gypsum by neutralising sulfuric acid in theliquid phase of the hydrolysis step in the conventional sulfate process.However, the gypsum product includes levels of impurities that reducethe market value of the gypsum. The liquid phase produced in hydrolysisstep (d) also includes sulfuric acid that can be neutralised to producegypsum. However, advantageously, this liquid phase is relatively free ofcontaminants because the titanyl sulfate precipitation step does notrecover substantial amounts (if any) of species (such as iron, chromium,manganese, and niobium) that are in solution in the leach liquor thatcould act as contaminants. Therefore, gypsum produced from this leachliquor is relatively pure.

Preferably the process includes separating a bleed stream of the leachliquor to minimise the build-up of species (such as vanadium, chromium,and niobium) in solution in the leach liquor.

The above-described process may be carried out as a continuous processor as a batch process.

Preferably the titaniferous material is ilmenite or altered ilmenite.

According to the present invention there is also provided hydratedtitanium oxides that have been produced by leaching a titaniferousmaterial (such as ilmenite) with sulfuric acid and forming a leachliquor that includes an acidic solution of titanyl sulfate and ironsulfate and thereafter hydrolysing titanyl sulfate and is characterisedin that the hydrated titanium oxides include coarse particles of atleast 0.005 mm (5 micron).

The process of the present invention includes the following typicalreactions.

Leaching:

FeTiO₃+2H₂SO₄→FeSO₄+TiOSO₄+2H₂O

Ferric reduction:

Fe₂(SO₄)₃+Fe°→3FeSO₄

Ferrous sulfate crystallisation:

FeSO₄+7H₂O→FeSO₄.7H₂O

Solvent extraction loading:

Ti(SO₄)₂+H₂O+R₃P═O→R₃P═O.TiOSO₄+H₂SO₄

Solvent extraction strip:

R₃P═O.TiOSO₄→R₃P═O+TiOSO₄

Hydrolysis:

TiOSO₄+2H₂O→TiO(OH)₂+H₂SO₄

Calcination:

TiO(OH)₂→TiO₂+H₂O

The improved sulfate process of the present invention is describedfurther with reference to the accompanying drawings, of which:

FIG. 1 is a flow sheet that illustrates one embodiment of the process ofthe invention; and

FIG. 2 is a flow sheet that illustrates one embodiment of the process ofthe invention.

With reference to the flow sheet of FIG. 1, in a Stage 1 Leach stepilmenite, leach liquor containing between 400 and 700 g/l sulfuric acidfrom a Stage 2 Leach step, and a reductant in the form of scrap iron aresupplied to a digester 3. The process operates on a continuous basiswith the feed materials being supplied continuously to the digester 3and reacted and unreacted materials being discharged continuously fromthe digester 3.

The Stage 1 Leach step solubilises a substantial component of theilmenite supplied to the digester 3 and produces a leach liquor thatcontains titanyl sulfate and iron sulfate. Typically, at the end of theleach the each liquor contains 20-100 and preferably 40-50 g/l titaniumand 50-100 g/l iron.

The leach liquor and partially and unreacted ilmenite that aredischarged continuously from the digester 3 are subjected to asolid/liquid separation step.

The solid phase from the solid/liquid separation step, which containsunreacted and partially reacted ilmenite, is transferred to the Stage 2Leach step. The Stage 2 Leach step is discussed further below.

The leach liquor from the solid/liquid separation step is transferredvia a heat exchanger 5 a to an iron sulfate crystallisation reactor 7.

The heat exchanger 5 a cools the leach liquor from a temperature of theorder of 110° C. to 60° C. The heat extracted by the heat exchanger 5 ais used elsewhere in the process, as discussed further below.

The leach liquor is cooled further, typically to 10-30° C. in the ironsulfate crystallisation reactor 7. Cooling the leach liquor precipitatesiron sulfate from the leach liquor in the iron sulfate crystallisationreactor 7. Typically, the crystallisation step reduces the concentrationof iron in the leach liquor to 40-50 g/l.

The leach liquor containing precipitated iron sulfate that is dischargedfrom the crystallisation reactor 7 is subjected to a furthersolid/liquid separation step which separates the precipitated ironsulfate from the leach liquor.

The solid phase from the solid/liquid separation step contains ironsulfate. The solid phase may also contain some species such as iron,manganese and aluminium. The solid phase is a by-product of the process.

The leach liquor from the solid/liquid separation step is transferredvia a heat exchanger 5 b to a solvent extraction reactor 9 and contactsa suitable solvent that extracts titanyl sulfate from the leach liquor.Typically, the leach liquor from the solid/liquid separation step is ata temperature of the order of 30° C. and the heat exchanger 5 b heatsthe leach liquor to a higher temperature, typically 50° C. Conveniently,the heat input for heat exchanger 5 b is heat recovered from the leachliquor by heat exchanger 5 a.

Suitable solvents are disclosed in Solex U.S. Pat. No. 5,277,816. Thesolvents include trioctylphosphine oxide and butyl dibutylphosphonate.The present invention is not confined to these extractants.

The solvent is used in conjunction with a modifier in the solventextraction step. Suitable modifiers include methyl isobutyl ketone(MIBK), di-isobutyl ketone (DIBK) and isotridecanol (ITA).

The solvent/titanyl sulfate mixture is separated from the leach liquor,and thereafter the titanyl sulfate is stripped from the solvent bywater.

The recovered solvent is returned to the solvent extraction reactor 9.

The resultant aqueous solution of titanyl sulfate, which typicallyincludes 10-100 g/l titanium in solution and 50-200 g/l sulfuric acid,is transferred to an hydrolysis reactor 11.

If the objective of the process is to produce feed material for pigmentproduction, the aqueous solution of titanyl sulfate may be processed inthe hydrolysis reactor 11 by conventional hydrolysis options such as theBlumenfeld and Mecklenberg processes.

If the objective of the process is to produce coarser feed material thanthat required for pigment production, the aqueous solution of titanylsulfate is processed in the hydrolysis reactor 11 as describedhereinafter.

Specifically, at start-up, the reactor 11 contains a starting solutionof sulfuric acid and solids. Typically, the solution contains 10-200 g/lacid and solids density of 10-200 g/l.

The titanyl sulfate solution is added at a controlled rate to thestarting solution. The addition of the solution results in the reactorfilling up to capacity and thereafter overflowing, whereafter the rateof overflow from the reactor 11 matches the rate of supply of titanylsulfate solution.

In the reactor 11 the sulfate ions in the titanyl sulfate solution aredisplaced by hydroxyl ions, with the result that hydrated titaniumoxides precipitate from the solution.

The solids in the starting solution act as seed for precipitation.Typically, the solids are hydrated titanium oxide or titanium dioxideparticles.

Typically, the residence time of titanyl sulfate solution in the reactor11 varies between 3 and 12 hours.

Subject to temperature and time conditions and control of solutionconcentration, there is controlled crystal growth in the hydrolysisreactor 11. Controlled crystal growth provides an opportunity to producetitania that ranges from fine to coarse particle sizes. In particular,controlled crystal growth provides an opportunity to produce coarsetitania of greater than 0.005 mm (5 micron) which can be used by way ofexample in the electrochemical reduction of titania to produce titanium.One important parameter for controlling crystal growth is theconcentration of titanium in solution within reactor 11. Specifically,it is preferred that the concentration be relatively low, of the orderof 10 g/l, within reactor 11 to achieve growth rather than nucleation oftitanium oxide particles.

The hydrolysis reactor 11 may be operated in batch mode. Morepreferably, the reactor is operated in continuous mode.

Moreover, if required, make-up water and solids can be added to thereactor 11.

In either the conventional pigment production hydrolysis or the abovecoarse particle size hydrolysis, the overflow from the reactor 11 iscollected as the product of the reactor 11.

The product from the hydrolysis reactor 11 is subjected to asolid/liquid separation step, which is facilitated by providing washwater.

The solid phase from the solid/liquid separation step, which containshydrated titanium oxides, is transferred to a calciner (not shown) andis calcined to produce titania. Depending on the circumstances, thesolid phase may be calcined at 1000° C. to produce titania.

In view of the efficiency of the solvent extraction step in confiningextraction substantially to titanium compounds, typically, the processproduces titania of very high purity, ie at least 99 wt. %.

Part or all of the liquid phase from the solid/liquid separation stepmay be reused in the process, for example as a source of acid in theStage 2 Leach step and/or as a source of water in washing steps in theprocess, as permitted by the overall water balance.

Alternatively, the liquid phase from the solid/liquid separation step,which contains sulfuric acid, is neutralised with lime and/or limestoneand thereby produces a gypsum product. In view of the efficiency of thesolvent extraction step in confining extraction to titanium compounds,the liquid phase contains minimal levels of contaminants (such as iron,vanadium and chromium) and therefore the gypsum is “clean” gypsum thatis commercially valuable in applications (such as the manufacture ofcement).

The raffinate from the solvent extraction step 9 contains relativelyhigh levels of sulfuric acid (250-700 g/l). The raffinate is transferredto the above-mentioned Stage 2 Leach step and is used as a leach liquor.In effect, the solvent extraction step recovers sulfuric acid and theacid can be used productively in the process. This enables a substantialreduction in waste when compared with the conventional sulfate process.In addition, the use of the raffinate as part of the acid feed for theprocess reduces the amount of fresh acid that is required in theprocess.

The Stage 2 Leach step is carried out in a digester 13.

The raffinate, and make-up concentrated sulfuric acid that is alsosupplied to the digester 13, leach the unreacted and partially reactedilmenite from the Stage 1 Leach and solubilise approximately 50% of theremaining ilmenite.

The product from the Stage 2 Leach is subjected to a solid/liquidseparation step.

The leach liquor from the solid/liquid separation step, which typicallycontains 400-700 g/l sulfuric acid, is transferred to the Stage 1 Leach,as mentioned above.

The solid phase from the solid/liquid separation step is substantiallymade up of silicate residue, and is a waste product of the process.

Make-up acid is required for the process since there are acid losses inthe separation of iron sulfate from the leach liquor and in theextraction of titanyl sulfate in the solvent extraction step.

The make-up acid may be added at any point in the flow sheet.

The addition of the acid in the Stage 2 Leach step is a preferredaddition point because it is thought that the introduction ofconcentrated acid at this point optimises the opportunity to leachilmenite, and it is beneficial to maintaining an efficient heat balance.

The flow sheet of FIG. 2 is very similar to that shown in FIG. 1 and thesame reference numerals are used to describe the same features in bothflow sheets.

The main difference between the flow sheets is that, whilst the FIG. 1flow sheet describes that the raffinate from the solvent extraction step9 is transferred to the Stage 2 Leach step and is used as a leachsolution in that step, in the FIG. 2 flow sheet the raffinate from thesolvent extraction step 9 is split into 2 separate streams and istransferred via the separate streams to the Stage 1 Leach step and theStage 2 Leach step, respectively, and is used as a leach solution inboth steps. In addition, whilst the FIG. 1 flow sheet describes that theliquid phase of the product from the Stage 2 Leach step is transferredto the Stage 1 Leach step, in the FIG. 2 embodiment the liquid phase istransferred to the leach liquor produced in the Stage 1 Leach step.

The applicant has carried out experimental work on a laboratory scaleand a pilot plant scale in relation to the above-described process.

In summary, the applicant has made the following findings in theexperimental scale work.

-   -   Fast leaching rates were achieved by leaching ilmenite in the        presence of an accelerant, such as scrap iron, sodium        thiosulfate, and sulfur dioxide.    -   Leach liquors containing up to 100 g/l titanium were produced.    -   The solvent extraction step resulted in a substantial upgrade in        purity of titania that was ultimately produced from the titanyl        sulfate extracted in the solvent extraction step.    -   The liquor stripped from the solvent in the solvent extraction        step contained high levels (at least 30 g/l) titanyl sulfate.    -   Raffinate can be used to leach ilmenite in the initial and the        further leach steps with or without make-up acid.    -   Two stage leaching is an effective leaching option, and the two        (or more than two) stage leaching can be carried out in a single        vessel with return of residual solid phase to the vessel and        addition of fresh ilmenite or in multiple vessels with the        residual solid phase produced in a 1^(st) vessel being supplied        to one or more than one other vessel.    -   There is a leach window (that is dependent on conditions such as        acid concentration, leach temperature, and leach time and        factors such as titanium ion concentration) in which it is        possible to avoid premature hydrolysis of hydrated titanium        oxides and premature precipitation of titanyl sulfate.

The laboratory scale and pilot plant scale work included leachingsamples of heavy mineral sands concentrates containing >50% ilmenite.

The leaching work included leaching work on a batch basis in 2 stages atatmospheric pressure with 30-50% w/w sulfuric acid at 95-120° C. for 3-5hours in each stage, and with additions of accelerant/reductant in theform of iron, sodium thiosulfate and sulfur dioxide in each stage.

The above leaching work was carried out with initial solids loadings of500 g/l and 200 g/l.

Table 1 is a summary of results of the above leaching work. TABLE 1Solids loading (in Stage 1) After Stage 1 After Stage 2 500 g/l 72% 87%200 g/l 63% 82%

Table 1 indicates that 2 stage leaching, under the conditions describedabove, is an effective leaching option.

The laboratory scale and the pilot plant scale work also includedsolvent extraction tests on leached ilmenite samples using a range ofsolvent extraction reagents and modifiers, including reagents of thetype disclosed in the U.S. Pat. No. 5,277,816 in the name of SolexResearch Corporation of Japan.

The solvent extraction tests were carried out after crystallisation ofexcess iron sulfate.

The reagents included, by way of example, Cyanex 923 [(C₈H₁₇)₃POequivalent] and the aliphatic diluent Shellsol D100A. The modifiersincluded, by way of example, methyl isobutyl ketone (MIBK), di-isobutylketone (DIBK) and isotridecanol (ITA).

Table 2 provides the composition of the feed solution and Table 3provides titanium enrichment factors in the loaded organic. TABLE 2 Ti84 g/l Ni 28 ppm Fe 66 g/l Si  8 ppm Mn 2.2 g/l Ca 42 ppm Cr 87 ppm Mg300 ppm  V 270 ppm Zn 66 ppm

TABLE 3 C923 Mix 1 Mix 2 Mix 3 Mix 4 Ti:Fe 275 450 407 909 1636 Ti:Mn˜inf. ˜inf. ˜inf. ˜inf. ˜inf. Ti:Cr ˜inf. ˜inf. ˜inf. ˜inf. ˜inf. Ti:V4.2 6.0 4.1 4.9 6.1 Ti:Ni 1.4 ˜inf. ˜inf. ˜inf. ˜inf. Ti:Si 0.14 ˜inf.˜inf. ˜inf. ˜inf. Ti:Ca ˜inf. 0.7 0.1 ˜inf. ˜inf. Ti:Mg 11 ˜inf. ˜inf.˜inf. ˜inf. Ti:Zn 0.3 ˜inf. 1.4 2.9 1.3 Ti (g/l) 9.1 15.0 14.0 20.0 9.0

Table 2 indicates that solvent extraction, under the conditionsdescribed above, is an effective means of separating titanium (in theform of titanyl sulfate) from contaminants.

The above solvent extraction tests also indicated that solventextraction is far more effective if a modifier is present. The modifierdid not appear to have any effect on the degree of extraction oftitanium. However, the modifier appeared to prevent the formation of anundesirable titanium-loaded phase that is not soluble in the diluent.Thus, without the modifier, only relatively dilute solutions of titaniumare possible.

The following Examples illustrate further the laboratory scale and pilotplant scale work carried out by the applicant.

EXAMPLE 1 Batch 1^(st) Stage Leach at Constant Acidity

1000 mL of raffinate containing 402 g/l free H₂SO₄, 24.6 g/l Fe²⁺, 2.0g/l Fe³⁺ and 3.3 g/l Ti was preheated to 110° C., in a glass reactorequipped with baffles and a Teflon agitator. 400 g of ilmenite,containing 30.4% Ti and 34.3% Fe and ground to 50% passing 32 μm, wasadded to this solution with sufficient agitation to fully suspend thesolids. A 6 mm mild steel rod was immersed into the slurry at a rate of0.5 cm/hour. Leaching was carried out for 6 hours. Aliquots of 98%sulfuric acid were added throughout to control the free acidity to 400g/l. After 6 hours a sample was withdrawn and filtered. Analysis of thesolution showed it to contain 397 g/l free H₂SO₄, 72.6 g/l Fe²⁺, 3.0 g/lFe³⁺ and 28 g/l Ti. The slurry was filtered, and the solids washed withwater and dried. 252.2 g of residue were obtained in this way,containing 31.9% Ti and 32.7% Fe.

EXAMPLE 2 Batch Two Stage Leach at Constant Acidity

1000 mL of synthetic raffinate containing 402 g/l free H₂SO₄ waspreheated to 105° C., in a glass reactor equipped with baffles and aTeflon agitator. 400 g of ilmenite, containing 30.4% Ti and 34.3% Fe andground to 50% passing 32 μm, was added to this solution with sufficientagitation to fully suspend the solids. 30 g of iron filings was added.Leaching was carried out for 5 hours. Aliquots of 98% sulfuric acid wereadded throughout to control the free acidity to 400 g/l. After 5 hours asample was withdrawn and filtered. Analysis of the solution showed it tocontain 387 g/l free H₂SO₄, 89.4 g/l Fe²⁺, 0.4 g/l Fe³⁺ and 48 g/l Ti.Heat and agitation were switched off and the slurry allowed to settleovernight. 750 mL of the clarified solution was removed and replacedwith an equal volume of fresh synthetic raffinate. Heat and agitationwere reinstated, and 30 g of iron filings were added. Leaching wascontinued at 110° C. for 5 hours. Aliquots of 98% sulfuric acid wereadded throughout to control the free acidity to 400 g/l. After 5 hours asample was withdrawn and filtered. Analysis of the solution showed it tocontain 373 g/l free H₂SO₄, 106 g/l Fe²⁺, 0.2 g/l Fe³⁺ and 38 g/l Ti.The slurry was filtered, and the solids washed with water and dried.57.5 g of residue were obtained in this way, containing 33.0% Ti and23.7% Fe.

EXAMPLE 3 Batch 1^(st) Stage Leach with Reducing Acidity

1000 mL of acidified raffinate containing 598 g/l free H₂SO₄, 31.3 g/lFe²⁺, 2.4 g/l Fe³⁺ and 9.2 g/l Ti was preheated to 110° C., in a glassreactor equipped with baffles and a Teflon agitator. 400 g of ilmenite,containing 30.4% Ti and 34.3% Fe and ground to 50% passing 32 μm, wasadded to this solution with sufficient agitation to fully suspend thesolids. A 6 mm mild steel rod was immersed into the slurry at a rate of0.5 cm/hour. Leaching was carried out for 6 hours. After 6 hours asample was withdrawn and filtered. Analysis of the solution showed it tocontain 441 g/l free H₂SO₄, 73.7 g/l Fe²⁺, 13.0 g/l Fe³⁺ and 47 g/l Ti.The slurry was filtered, and the solids washed with water and dried.223.6 g of residue were obtained in this way, containing 32.0% Ti and32.8% Fe.

EXAMPLE 4 Batch 2^(nd) Stage Leach with Reducing Acidity

1000 mL of synthetic raffinate containing 593 g/l free H₂SO₄, waspreheated to 105° C., in a glass reactor equipped with baffles and aTeflon agitator. 400 g of 1^(st) stage leach residue, containing 32.0%Ti and 31.3% Fe was added to this solution with sufficient agitation tofully suspend the solids. A 6 mm mild steel rod was immersed into theslurry at a rate of 0.5 cm/hour. Leaching was carried out for 6 hours.After 6 hours a sample was withdrawn and filtered. Analysis of thesolution showed it to contain 476 g/l free H₂SO₄, 29.0 g/l Fe²⁺, 10.4g/l Fe³⁺ and 32.5 g/l Ti. The slurry was filtered, and the solids washedwith water and dried. 267 g of residue were obtained in this way,containing 31.9% Ti and 30.7% Fe.

EXAMPLE 5 Pilot Plant 1^(st) Stage Leach with Reducing Acidity

39 L of 98% sulfuric acid was added to 243 L of raffinate containing 358g/l free H₂SO₄ and 7 g/l Ti, in a fibre reinforced plastic (FRP) tank of300 L capacity, equipped with a FRP axial turbine. The resultingsolution, which contained 579 g/l free acid, 27.9 g/l Fe²⁺ and 5.6 g/lFe³⁺, was preheated to 95° C. 116 kg of unground ilmenite, containing31.1% Ti and 34.1% Fe was added to this solution with sufficientagitation to fully suspend the solids. A group of ten 10 mm mild steelrods of length 29 cm was immersed into the slurry. Leaching was carriedout for 6 hours at 105° C. The slurry was filtered using a pressurefilter, to produce approximately 260 L of solution. Analysis of thesolution showed it to contain 461 g/l free H₂SO₄, 72.6 g/l Fe²⁺, 9.0 g/lFe³⁺ and 41 g/l Ti.

EXAMPLE 6 Pilot Plant 1^(st) Stage Leach with Constant Acidity

A single stage leach pilot plant was assembled, consisting of 5 stirredFRP tanks of 10 L capacity each, equipped with FRP double axialturbines, and silica jacketed electric immersion heaters. Ilmeniteground to 50% passing 32 μm was fed to the first tank at 750 g/hourusing a screw feeder. SX pilot plant raffinate of composition 404 g/lfree H₂SO₄, 36.1 g/l Fe²⁺, 3.2 g/l Fe³⁺ and 10 g/l Ti, was also pumpedinto the first tank at a rate of 62.5 mL/min. The temperature wasmaintained at 110° C. in all tanks. 98% sulfuric acid was added to thefirst two tanks to control the acidity to 400 g/l. Mild steel rods ofdiameter 10 mm were inserted into each tank at a rate of 1 cm/hr. Slurrywas thence allowed to flow by gravity to a FRP thickener equipped withFRP rakes. Thickener overflow-solution and underflow slurry werecollected and stored. The pilot plant was operated continuously for 92hours. During the final 48 hours of operation the average composition ofthe solution in each tank was as set out below in Table 4. TABLE 4Continuous Pilot Plant 1^(st) Stage Leach Results Free Acid g/l Fe²⁺ g/lFe³⁺ G/l Ti g/l Tank 1 411 48 2.7 16 Tank 2 404 56 2.5 20 Tank 3 402 652.4 29 Tank 4 395 68 4.2 36 Tank 5 391 65 6.1 34 Thickener 388 70 3.0 33overflow

EXAMPLE 7 SX Bench Tests with Counter-Current Extraction

Three groups of counter-current bench tests were carried out to simulatethe SX extraction circuit of the pilot plant operation. Each groupinvolved 5 cycles and the data indicated that a steady state wasachieved. The organic phase contained 30% vol Cyanex 923 as theextractant, 5% vol DIBK as the modifier and 65% vol Shellsol D100A asthe diluent. At the O/A ratio of 2, 3 and 4, the organic loading was 16,11 and 8 g/l Ti; the extraction efficiency was 97.8, 99.7 and 99.9%; thetitanium concentrations of raffinate were 450-910, 80-120 and 24-28 mg/lrespectively. The separation between Ti and Cr, Mg, Mn, Ni approachedperfect with the loaded organic containing 0 mg/l of Cr, Mg, Mn and Ni.The test was carried out on flask shaker in an incubator. The major testconditions are shown as follows: Temperature: 50° C Mixing time: 45-60minutes Settling time: 15 minutes O/A ratio: 2, 3, 4

The results are summarised in Table 5. TABLE 5 Composition of Organic,Feed and Raffinate at Given O/A Ratio (mg/l) H₂SO₄ Ca Cr Fe Mg Mn Ni PSi Ti Y Zn Feed 441000 10 52 39000 220 1400 7 0 27 32000 120 47 O/A Org28000 4 0 708 0 0 0 20200 6 15600 13 0 2 Raf 327000 14 55 37200 222 15009 0 76 676 108 47 O/A Org 23000 1 0 1920 0 0 0 20200 5 11000 12 6 3 Raf314000 14 55 34800 224 1540 9 4 3 89 106 48 O/A Org 22000 2 1 2160 0 0 120000 4 8360 10 0 4 Raf 296000 15 58 34600 240 1600 10 7 6 25 100 51

EXAMPLE 8 SX Bench Stripping Test

The loaded organic that contained 30% vol Cyanex 923 as the extractant,5% vol DIBK as the modifier and 65% vol Shellsol D100A as the diluentwas stripped with water with various O/A ratio. The test was carried outusing a flask shaker in an incubator. The major test conditions areshown as follows: Temperature: 50° C. Mixing time: 60 minutes Settlingtime: 20 minutes O/A ratio: 1/3, 1/1, 3/1, 5:/1, 10/1, 20/1 and 30/1

The results are summarised in Table 6. TABLE 6 Composition of Organicand Stripping Solution at Given O/A Ratio (mg/l) H₂SO₄ Ca Cr Fe Mg Mn NiP Si Ti Y Zn Loaded Organic n.a 13 0 44 3 0 0 20000 4 18000 12 8 O/A Orgn.a 14 0 47 3 0 0 21000 5 3400 8 19 1/3 Strip* 14700 2 0 16 0 0 0 6 03800 3 0 O/A Org n.a 14 0 18 0 0 0 20000 6 11000 12 10 1/1 Strip 44100 237 0 0 0 0 7 0 7200 6 1 O/A Org n.a 12 0 16 0 0 0 20000 5 15000 13 9 3/1Strip 103000 2 0 97 0 0 0 7 0 8000 7 2 O/A Org n.a 13 0 30 4 0 0 20000 416000 17 25 5/1 Strip 156900 0 0 160 0 0 0 3 0 7200 7 0 O/A Org n.a 3 026 3 0 0 20000 0 17000 18 20 10/1  Strip 225600 2 0 280 0 0 0 0 0 5800 63 O/A Org n.a 0 0 40 2 0 0 20000 0 17000 17 20 20/1  Strip 343000 4 0490 0 0 0 0 1 4700 8 0 O/A Org n.a 10 0 41 0 0 0 19000 0 17000 15 630/1  Strip 392300 2 1 710 0 0 0 0 0 4500 7 2*White precipitation formed in both organic and aqueous phase

EXAMPLE 9 SX Pilot Plant Operation

The pilot plant operation was carried out with a device that involvedtwo extraction cells, one scrub cell and four stripping cells. Theeffective volume of mixer and settler of each cell were 1.675 and 8.000liter respectively. The stripping involved two stages: the lead stripwith hydrolysis thickener underflow overflow and lag stripping withwater respectively. The major operational conditional were as follows:Temperature: 50° C. Organic Composition: 25% v/v Cyanex 923, 5% v/vIso-tri-decanol and 70% v/v ShellSol D100A. The capacity of organic was15.7 g/l Ti. Feed Composition: 36 g/l Ti, 410 g/l H₂SO₄, 47 g/l Feincluding 4.0 g/l Fe³⁺ Mixing time: 5-10 minutes Settling time: 40minutes O/A flow ratio of extraction: ˜5:1. Organic flow: 165 mL/min;Feed flow: 33 mL/min O/A flow ratio of scrub: ˜10:1. Org. flow: 165mL/min; Aqu. flow: 17 mL/min O/A flow ratio of lead strip: ˜4:1. Org.flow: 165 mL/min; Aqu. flow: 41 mL/min O/A flow ratio of lag strip:˜8:1. Org. flow: 165 mL/min; Aqu. Flow: 21 mL/min

The results are summarised in Table 7. TABLE 7 Average Compositions ofAqueous and Organic Liquor Tot. H₂SO₄ Ca Cr Fe Fe³⁺ Mg Mn Ni P Si Ti VZn g/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l Feed410 58 43 47000 4000 430 1900 4 22 17 36000 200 41 Raff 345 48 29 305295700 319 1371 4 2 10 5324 122 30 Lead Strip 178 27 1 465 465 2 0 0 13 511306 16 1 Lag Strip 91 28 2 224 224 3 1 0 15 5 16329 21 1 Loaded Org.(E2) n.a 3 1 160 160 1 0 0 16647 0 15059 22 2 Strip Org. (ST1) n.a 3 023 23 I 0 0 17000 0 7224 11 3

EXAMPLE 10 Pilot Plant Operation

The pilot plant operation was carried out with a device that involvedtwo extraction cells, one scrub cell and four stripping cells. Theeffective volume of mixer and settler of each cell were 1.675 and 8.000liter respectively. The stripping involved two stages: the lead stripwith hydrolysis thickener overflow and lag stripping with 50 g/l H₂SO₄respectively. The major operational conditional were as follows:Temperature: 50° C. Organic Composition: 25% v/v Cyanex 923, 5% v/vIso-tri-decanol and 70% v/v ShellSol D100A. The capacity of organic was15.7 g/l Ti. Feed Composition: 36 g/l Ti, 410 g/l H₂SO₄, 47 g/l Feincluding 4.0 g/l Fe³⁺ Mixing time: 5-10 minutes Settling time: 40minutes O/A flow ratio of extraction: ˜5:1. Organic flow: 165 mL/min;Feed flow: 32 mL/min O/A flow ratio of scrub: ˜10:1. Org. flow: 165mL/min; Aqu. flow: 17 mL/min O/A flow ratio of lead strip: ˜4:1. Org.flow: 165 mL/min; Aqu. flow: 41 mL/min O/A flow ratio of lag strip:˜8:1. Org. flow: 165 mL/min; Aqu. Flow: 21 mL/min

The results are summarised in Table 8. TABLE 8 Average Compositions ofAqueous and Organic Liquor Tot. H₂SO₄ Ca Cr Fe Fe³⁺ Mg Mn Ni P Si Ti VZn g/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l mg/l Feed415 83 45 49000 5200 450 1800 5 17 21 34000 350 30 Raff 363 65 29 295005000 340 1258 5 0 11 3950 115 29 Lead Strip 223 53 2 408 408 5 0 0 24 717167 28 1 Lag Strip 113 54 1 56 56 6 0 0 7 7 9300 14 1 Loaded Org. n.a15 1 177 177 0 0 0 19167 2 15917 23 3 (E2) Stripped n.a 2 0 0 0 0 0 019750 1 10358 17 0 Org. (ST1)

EXAMPLE 11 Batch Hydrolysis

1000 mL of SX pilot plant loaded strip liquor containing 123 g/l freeH₂SO₄, 0 g/l Fe²⁺, 0.26 g/l Fe³⁺ and 12 g/l Ti was pretreated with 1 gof aluminium foil overnight at room temperature. Titration withdichromate with sodium diphenylamine sulfonate as indicator showed theresulting solution to contain 2.4 g/l Ti³⁺. 500 mL of water containing100 g/l free H₂SO₄, and 0.5 g of TiO(OH)₂ seed, was preheated to 95° C.,in a glass reactor equipped with baffles and a Teflon agitator. Thetreated loaded strip liquor was then pumped into the reactor at 2.8mL/min over 6 hours. The reaction mixture was allowed to stir for afurther 30 minutes then a sample was withdrawn and filtered. Analysis ofthe solution showed it to contain 147 g/l free H₂SO₄, 0.24 g/l Fe and2.3 g/l Ti. The slurry was filtered, and the solids washed with waterand dried. Filtration was found to be very fast. 22.6 g of residue wereobtained in this way, containing 45.0% Ti, 3.9% S and <0.02% Fe with d₅₀particle size approximately 8.5 micron.

EXAMPLE 12 Pilot Plant Hydrolysis

A single stage hydrolysis pilot plant was assembled, consisting of 2stirred FRP tanks of 5 L capacity each, equipped with FRP double axialturbines, and silica jacketed electric immersion heaters. SX pilot plantloaded strip liquor containing 206 g/l free H₂SO₄, 0 g/l Fe²⁺, 0.2 g/lFe³⁺ and 25 g/l Ti, was pumped into the first tank at a rate of 10mL/min. The temperature was maintained at 95° C. in each tank. Water wasadded to the first tank to control the acidity to 140 g/l, requiring aflow of 8.5 mL/min. Additional water was added to the second tank at 5mL/min to control the acidity to 100 g/l. Slurry was thence allowed toflow by gravity to a FRP thickener equipped with FRP rakes. Thickeneroverflow solution was collected and stored. The thickener underflowslurry were was collected and filtered by vacuum filtration. Filtrationof the underflow was found to be very fast. The d₅₀ particle size wasfound to be 7.2 micron. The pilot plant was operated continuously for 42hours. During the final 30 hours of operation the average composition ofthe solution in each tank was as follows: TABLE 9 Continuous Pilot PlantHydrolysis Results Free Acid g/l Fe g/l Ti g/l Tank 1 136 0.13 4.6 Tank2 105 0.09 1.0 Thickener overflow 109 0.09 1.0

EXAMPLE 13 Laboratory Scale Calcination

A 2.6 g sample of dried TiO(OH)₂ produced according to Example 11 wascalcined in an alumina crucible, using a muffle furnace at 1000° C. for1 hour. On removal from the furnace the cooled calcine was found by XRFto contain 59.8% Ti, 0.07% Fe, <0.02% S and less than detection limitfor Si, Al, Mn, Mg, Cr, V and Zn.

Many modifications may be made to the process of the present inventiondescribed above without departing from the spirit and scope of thepresent invention.

By way of example, whilst the above-described flow sheet describes thatthe Stage 1 and Stage 2 Leach steps are carried out in single digesters3 and 13, respectively, the present invention is not so limited andextends to arrangements that include multiple digesters for each stage.

In addition, whilst the above-described flow sheet describes that theStage 1 and Stage 2 Leach steps are carried out in separate digesters 3and 13, respectively, the present invention is not so limited andextends to arrangements in which leaching of titaniferous material iscarried out in a single digester, with return of residual solid phase tothe digester and direct supply of raffinate from the solvent extractionstep 9 to the digester.

1. A sulfate process for producing titania from a titaniferous materialwhich includes the steps of: (a) leaching the solid titaniferousmaterial with a leach solution containing sulfuric acid and forming aleach liquor that includes an acidic solution of titanyl sulfate(TiOSO₄) and iron sulfate (FeSO₄); (b) separating the leach liquor and aresidual solid phase from the leach step (a); (c) separating titanylsulfate from the leach liquor from step (b); (d) hydrolysing theseparated titanyl sulfate and forming a solid phase containing hydratedtitanium oxides; (e) separating the solid phase containing hydratedtitanium oxides and a liquid phase that are produced in the hydrolysisstep (d); and (f) calcining the solid phase from step (e) and formingtitania; (g) a further leach step of leaching the residual solid phasefrom step (b) with a leach solution containing sulfuric acid and forminga leach liquor that includes an acidic solution of titanyl sulfate andiron sulfate and a residual solid phase; (h) separating the leach liquorand the residual solid phase from step (g); and (i) supplying theseparated leach liquor to the leach step (a) and/or mixing the separatedleach liquor with the leach liquor from step (b).
 2. The process definedin claim 1 further comprising carrying out the leach step (a) and thefurther leach step (g) in the same vessel.
 3. The process defined inclaim 2 further comprising returning the residual solid phase from step(b) to the vessel, wherein the residual solid phase forms part of thetitaniferous material subjected to leaching in the leach step (a). 4.The process defined in claim 1 further comprising carrying out the leachstep (a) and the further leach step (g) in a separate vessel or vessels.5. The process defined in claim 4 wherein the further leach step (g)includes supplying the residual solid phase from step (b) to the vesselor vessels.
 6. The process defined in claim 1 wherein the leach step (a)and/or the further leach step includes selecting and/or controlling theleach conditions in the leach step or steps to avoid undesirable amountsof premature hydrolysis of hydrated titanium oxides and undesirableamounts of premature precipitation of titanyl sulfate.
 7. The processdefined in claim 6 wherein the leach conditions include any one or morethan one of acid concentration, leach temperature and leach time.
 8. Theprocess defined in claim 6 further comprising selecting and/orcontrolling the acid concentration to be at least 350 g/l sulfuric acidthroughout the leach step (a) and/or the further leach step (g) whenoperating at a leach temperature in the range of 95° C. to the boilingpoint in order to avoid premature hydrolysis.
 9. The process defined inclaim 6 further comprising selecting and/or controlling the acidconcentration at the end of the leach step (a) and/or the further leachstep (g) to be less than 450 g/l when operating at a leach temperaturein the range of 95° C. to the boiling point in order to avoid anundesirable amount of premature precipitation of titanyl sulfate. 10.The process defined in claim 6 further comprising selecting and/orcontrolling the leach conditions so that the titanium ion concentrationin the leach liquor is less than 50 g/l in the leach liquor at the endof the leach step (a) and/or the further leach step (g).
 11. The processdefined in claim 10 further comprising selecting and/or controlling theleach conditions so that the titanium ion concentration in the leachliquor is 40-50 g/l in the leach liquor at the end of the leach step (a)and/or the further leach step (g).
 12. The process defined in claim 1further comprising carrying out the leach step (a) and/or the furtherleach step (g) in the presence of an additive that accelerates the rateof leaching the titaniferous material.
 13. The process defined in claim12 wherein the leaching accelerant is selected from a group thatincludes iron, a titanium (III) salt, a thiosulfate salt, sulfur dioxideor any other reduced sulfur containing species.
 14. The process definedin claim 1 further comprising carrying out the leach step (a) and/or thefurther leach step (g) in the presence of a reductant that reducesferric ions to ferrous ions in the acidic solution or solutions oftitanyl sulfate and iron sulfate produced in the leach step (a).
 15. Theprocess defined in claim 14 wherein the reductant is selected from agroup that includes iron, a titanium (III) salt, a thiosulfate salt,sulfur dioxide or any other reduced sulfur containing species.
 16. Theprocess defined in claim 1 wherein the leach step (a) solubilises atleast 50% by weight of the titaniferous material supplied to the leachstep.
 17. The process defined in claim 1 further comprising the steps ofprecipitating iron sulfate from the leach liquor from step (b) andseparating precipitated iron sulfate from the leach liquor prior to thetitanyl sulfate separation step (c).
 18. The process defined in claim 1further comprising using at least part of the leach liquor remainingafter separation of titanyl sulfate in step (c) as at least part of theleach solution in the leach step (a) and/or in the further leach step(g).
 19. The process defined in claim 18 wherein the titanyl sulfateseparation step (c) includes a solvent extraction step of extractingtitanyl sulfate from the leach liquor from step (b) into a solvent andthereafter stripping titanyl sulfate from the solvent and forming asolution that contains titanyl sulfate and thereafter hydrolysing thetitanyl sulfate-containing solution in the hydrolysis step (d).
 20. Theprocess defined in claim 19 further comprising using at least part of araffinate from the solvent extraction step as at least part of the leachsolution in leach step (a) and/or in the further leach step (g).
 21. Theprocess defined in claim 20 wherein the leach solution in the leach step(a) and the further leach step (g) includes the raffinate and make-upfresh sulfuric acid.
 22. The process defined in claim 20 wherein theraffinate from the solvent extraction step has an acid concentration ofat least 250 g/l sulfuric acid.
 23. The process defined in claim 19wherein the solvent extraction step includes contacting the leach liquorwith the solvent which includes a modifier.
 24. The process defined inclaim 1 further comprising controlling the hydrolysis step (d) toproduce a selected particle size distribution of the hydrated titaniumoxides product.